Method for producing mixtures of 1,4-butanediol, tetrahydrofuran and γ-butyrolactone

ABSTRACT

In a process for preparing 1,4-butanediol, tetrahydrofuran and γ-butyrolactone by oxidation of butane to give a product stream comprising maleic anhydride, absorption of maleic anhydride from the product stream using a high-boiling alcohol to give a liquid absorption product comprising monoesters and diesters of maleic acid and also high-boiling alcohol, after-esterification of the liquid absorption product and subsequent hydrogenation of the after-esterified product in the liquid phase, the high-boiling alcohol is a polyhydric alcohol having a boiling point at atmospheric pressure of above 233° C. and the after-esterified product has an acid number of less than 30 mg KOH/g and a water content of less than 1% by weight.

This application is a 371 of PCT/EP99/02685 filed Apr. 21, 1999.

The present invention relates to a process for preparing mixturescomprising 1,4-butanediol, tetrahydrofuran (THF) and γ-butyrolactone(GBL) from the product gas stream from a reactor for the oxidation ofbutane by a) absorption of maleic anhydride (MA) from the product gasstream using a high-boiling alcohol, b) conversion of the maleicmonoester formed into the maleic diester and c) hydrogenation of thelatter in the liquid phase.

Numerous processes for converting MA into the corresponding monoestersand diesters and hydrogenating the latter are known.

EP-B 0 149 144, EP-B 0 206 194 and EP-B 0 212 121 describe processes forthe continuous separation of MA from gaseous reaction mixtures which areobtained in the catalytic oxidation of hydrocarbons. Here, theMA-containing, gaseous reaction mixture is brought into contact with amonohydric alcohol. The gaseous materials formed are brought intocontact with dicarboxylic diesters (EP-B 0 206 194 dibutyl fumarate orsuccinate, EP B 0 212 121 diesters of fumaric, succinic or maleic acid,EP-B 0 149 144 dibutyl maleate) in a countercurrent process and theliquid process product is taken off at the bottom. The liquid processproduct comprises predominantly the corresponding monoalkyl and dialkylesters of maleic acid. These are heated at from 110° C. to 200° C. tocomplete the esterification and then provide a suitable startingmaterial for the hydrogenation to form 1,4-butanediol.

A disadvantage of this process is that, apart from the alcohol, anadditional material (dicarboxylic diester) which transfers the gaseousreaction products formed from alcohol and MA into the liquid phase iscirculated. This must not be hydrogenated in a hydrogenation, so that itis always only possible to achieve incomplete conversion, which meanscomplicated control of the hydrogenation. Furthermore, the dibutyl esterof a dicarboxylic acid used as absorption medium has to be separatedfrom butanediol when the product is worked up by distillation, whichcomplicates the work-up.

DE-A 31 06 819, too, describes a process for preparing 1,4-butanediol bycatalytic hydrogenation of a mixture which is obtained by treatment ofMA-containing, gaseous reaction mixtures with aliphatic alcohols. Theabsorption of the MA is carried out using monohydric or dihydricalcohols having boiling points above 180° C. An additional absorptionmedium is not necessary. The subsequent after-esterification of theabsorption product is carried out at from 120° C. to 150° C. Thediester-containing stream is finally catalytically hydrogenated, but thespace-time yield (0.04 kg of butanediol/liter×hour) is low since theacid number after the after-esterification is too high.

Although the abovementioned processes can be carried out industrially,they have disadvantages which lead to high production costs. Thesedisadvantages are, in particular, complicated control of thehydrogenation and a low space-time yield in the catalytic hydrogenation.

It is an object of the present invention to provide a process whichrequires very little absorption medium, achieves a good space-time yieldin the catalytic hydrogenation using inexpensive catalysts and, in theseparation of the desired products, forms no mixtures which aredifficult to separate.

We have found that this object is achieved by a process for preparing1,4-butanediol, tetrahydrofuran (THF) and γ-butyrolactone (GBL) byoxidation of butane to give a product stream comprising maleicanhydride, absorption of maleic anhydride from the product stream usinga high-boiling alcohol to give a liquid absorption product comprisingmonoesters and diesters of maleic acid and also high-boiling alcohol,after-esterification of the liquid absorption product and subsequenthydrogenation of the after-esterified product in the liquid phase. Inthe process of the present invention, the high-boiling alcohol is apolyhydric alcohol having a boiling point at atmospheric pressure ofabove 233° C. and the after-esterified product has an acid number ofless than 30 mg KOH/g and a water content of less than 1% by weight.

In the context of the present invention, absorption means the separationof MA from a product stream obtained by oxidation of butane using ahigh-boiling alcohol to carry out this separation. In this absorptionstep, the MA reacts with the alcohol to give a maleic monoester whichconstitutes the main product in the absorption product.

In this process, a liquid process product is obtained directly byaddition of the esterifying alcohol under the prevailing reactionconditions. The additional use of a high-boiling absorption medium suchas a diester of a dicarboxylic acid is not necessary. By this means, acomplicated, later separation of the desired products is avoided. Owingto the low acid number, the hydrogenation can be carried out usinginexpensive catalysts and in good yields using the after-esterifiedproduct. This allows cost savings.

Catalytic oxidation of butane or another hydrocarbon in the gas phaseover a vanadium pentoxide catalyst activated with, for example, MoO₃gives a product stream comprising MA and by-products such as carbondioxide, acetic acid and acrylic acid. In the process of the presentinvention, this MA-containing, gaseous product stream is brought intocontact with a high-boiling alcohol to absorb the MA. In an advantageousembodiment of the invention, the reaction gas is fed in below thesurface of the liquid, high-boiling alcohol, e.g. through an immersedtube. Preference is given to a process in which the reaction mixture isintroduced directly from below into an absorption column in which theliquid, high-boiling alcohol flows toward the reaction mixture. Carryingout the MA absorption in a column or a plurality of columns connected inseries is preferred.

Suitable columns are, for example, bubble cap tray columns, columnscontaining units of packing or columns packed with loose packingelements, with preference being given to using the latter. These can beprovided with intermediate coolers in order to remove the heat ofabsorption.

The MA content of the product stream from the butane oxidation is notcritical for the process of the present invention. In customaryprocesses, it is in the range from 0.5 to 2% by volume. The temperatureof the product gas stream is likewise generally not critical. It should,if possible, be not below the dew points and melting points of theindividual components in order to avoid caking, i.e. it shouldpreferably be at least 100° C.

The by-products occurring in the oxidation of butane, for example CO₂,acetic acid and acrylic acid, are likewise not critical for the process.They are either mostly carried away with the waste gas stream or go intothe absorption medium. If they do not react with the esterifyingalcohol, they are removed in the thermal treatment of the liquidabsorption product (partly with the water).

Alcohols used in the process of the present invention are polyhydricalcohols having a boiling point at atmospheric pressure of above 233°C., preferably above 250° C. As polyhydric alcohols preference is givento using dihydric to tetrahydric alcohols, particularly preferablydihydric alcohols (diols). Examples of alcohols used are polyethyleneglycols, α, ω-diols and cyclohexanedimethanols, e.g. 1,5-pentanediol,1,6-hexanediol, 1,7-heptanediol, 1,8-octanediol, 1,12-dodecanediol,1,4-cyclohexanedimethanol, 1,3-cyclohexanedimethanol,trimethylolpropane, neopentyl glycol, triethylene glycol, tetraethyleneglycol and pentaethylene glycol, very particularly preferably1,6-hexanediol and 1,4-cyclohexanedimethanol. As trihydric alcohol, itis possible to use glycerol; as tetrahydric alcohol, pentaerythritol.The alcohols can be used in pure form or as mixtures of variousalcohols.

In general, the alcohol is used in a molar excess over the MA of up to30, preferably up to 15, very particularly preferably up to 5.

In the absorption, a monoester of maleic acid is first formed from theMA and the alcohol. This monoester is high-boiling and thus no longervolatile at the reaction temperatures. The absorption can be carried outso that the after-esterification of the monoester to form the diesteralso takes place in the apparatus used for the absorption of MA. In thisreaction, the monoester can react with further free alcohol or withanother monoester molecule. The water eliminated in the reaction can, ifthe temperatures are high enough, be removed with the waste gas stream.

If relatively large amounts of monoester are still present in the liquidproduct, preference is given to carrying out a thermalafter-esterification in a residence time vessel and thus completing theesterification to form the diester. The after-esterification can becarried out with or without addition of an esterification catalyst.

Esterification catalysts which can be used are, in principle, allhomogeneous or heterogeneous catalysts known for the esterification ofacids. Preference is given to heterogeneous catalysts such as TiO₂,ZrO₂, Al₂O₃, SiO₂, silicates, zeolites, heteropolyacids or acid ionexchangers.

Particular preference is given to carrying out the esterificationwithout catalyst, purely thermally. The monoester-containing stream isthen after-esterified at temperatures of from 160° C. to 300° C. ingeneral, preferably from 160° C. to 250° C., particularly preferablyfrom 180° C. to 240° C. The water of reaction is generally removedcontinuously by distillation. The residence time usually is a maximum of3 hours, in general from 0.1 to 3 hours, preferably from 0.2 to 2.5hours, particularly preferably from 0.3 to 2 hours. To remove the waterof reaction, it is also possible to carry out the after-esterificationin a stripping column and remove the water liberated as water vapor bymeans of stripping gas. It is also possible to apply a vacuum to make iteasier to remove the water of reaction. Another possibility is to addentrainers for water, e.g. aromatic hydrocarbons. The water contentafter the esterification is less than 1% by weight, preferably less than0.5%, particularly preferably less than 0.2%.

Owing to the choice of esterification conditions, the free acid contentof the esterification product is low. The free acid content (measured bytitration) before the hydrogenation is less than 30 mg KOH/g, preferablyless than 20 mg KOH/g, particularly preferably less than 10 mg KOH/g.

The hydrogenation can be carried out in the presence of an inert organicsolvent, particularly if only small quantities are to be hydrogenated.For example, the esterification product can be diluted with an inertorganic solvent, preferably ethylene glycol dimethyl ether, to make itmore pumpable. When carrying out the reaction industrially, anadditional solvent is generally dispensed with.

The hydrogenation is carried out batchwise or continuously in the liquidphase over fixed-bed or suspended or homogeneous, soluble catalysts.Preference is given to a continuous procedure. In the case of fixed-bedcatalysts, hydrogenation can be carried out in the downflow or upflowmode, with or without product recirculation. One or more reactors can beoperated in series or in parallel. For example, it is possible tohydrogenate predominantly the C—C double bond of the maleic esters inthe first reactor, forming succinic esters, and subsequently to continuethe hydrogenation in a second reactor to give butanediol, THF and GBL.

The hydrogenation is carried out generally at temperatures of from 70°C. to 350° C., preferably from 80° C. to 300° C., particularlypreferably from 80° C. to 260° C., and at pressures of from 20 bar to350 bar, preferably from 40 bar to 320 bar in general, particularlypreferably from 60 bar to 300 bar.

Hydrogenation catalysts which can be used in the process of theinvention are, in general, heterogeneous or homogeneous catalystssuitable for the hydrogenation of carbonyl groups. Preference is givento heterogeneous catalysts. Examples are described in Houben-Weyl,Methoden der Organischen Chemie, Volume IV/1c, pp. 16 to 26, GeorgThieme Verlag, 1980.

Among these hydrogenation catalysts, preference is given to those whichcomprise at least one element of groups Ib, VIb, VIIb and VIII, andIIIa, IVa and Va of the Periodic Table of the Elements, in particularcopper, chromium, rhenium, cobalt, rhodium, nickel, palladium, iron,platinum, indium, tin and antimony. Particular preference is given tocatalysts comprising copper, cobalt, palladium, platinum or rhenium.Among these, very particular preference is given to those which comprisecopper.

One example of a type of catalyst which can be used for the process ofthe present invention are unsupported catalysts. In these, thecatalytically active metals are present essentially without supportmaterials. Examples are the Raney catalysts, e.g. those based on Ni, Cuor cobalt. Other examples are Pd black, Pt black, Cu sponge, alloys ormixtures of, for example, Pd/Re, Pt/Re, Pd/Ni, Pd/Co or Pd/Re/Ag.

The catalysts used in the process of the present invention can also beprecipitated catalysts. Such catalysts can be prepared by precipitatingtheir catalytically active components from their salt solutions, inparticular from their nitrate and/or acetate solutions, for example byaddition of alkali metal hydroxide and/or alkaline earth metal hydroxideand/or carbonate solutions, e.g. as sparingly soluble hydroxides,hydrated oxides, basic salts or carbonates. The precipitates obtainedare subsequently dried and then converted into the corresponding oxides,mixed oxides and/or mixed-valence oxides by calcination at generallyfrom 300 to 700° C., preferably from 400 to 600° C. These oxides arereduced by treatment with hydrogen or hydrogen-containing gases atgenerally from 50 to 700° C., preferably from 100 to 400° C., to givethe corresponding metals and/or oxidic compounds of lower oxidationstates and thus converted into the actual catalytically active form. Thereduction is generally carried out until no more water is formed.

In the preparation of precipitated catalysts which comprise a supportmaterial, the precipitation of the catalytically active components canbe carried out in the presence of the support material concerned. Thecatalytically active components can also be advantageously precipitatedsimultaneously with the support material from the salt solutionsconcerned.

In the process of the present invention, preference is given to usinghydrogenation catalysts which comprise the metals or metal compoundswhich catalyze the hydrogenation deposited on a support material. Apartfrom the abovementioned precipitated catalysts, which comprise a supportmaterial in addition to the catalytically active components, supportmaterials in general in which the hydrogenation-active catalyticcomponents have been applied to a support material by, for example,impregnation are also suitable for the process of the present invention.

The method by which the catalytically active metals are applied to thesupport is generally not critical and the application can be carried outin various ways. The catalytically active metals can be applied to thesesupport materials by, for example, impregnation with solutions orsuspensions of the salts or oxides of the elements concerned, drying andsubsequent reduction of the metal compounds to the corresponding metalsor compounds of lower oxidation states by means of a reducing agent,preferably using hydrogen or complex hydrides.

Another possible way of applying the catalytically active metals to thesupports is to impregnate the supports with solutions of salts which arereadily decomposed thermally, e.g. nitrates, or complexes which arereadily decomposed thermally, e.g. carbonyl or hydrido complexes, of thecatalytically active metals and to heat the support which has beenimpregnated in this way at from 300 to 600° C. so as to effect thermaldecomposition of the adsorbed metal compounds. This thermaldecomposition is preferably carried out under a protective gasatmosphere. Suitable protective gases are, for example, nitrogen, carbondioxide, hydrogen or the noble gases.

The catalytically active metals can also be deposited on the catalystsupport by vapor deposition or by flame spraying.

The content of the catalytically active metals in the supportedcatalysts is in principle not critical for the success of the process ofthe present invention. Higher contents of catalytically active metals inthe supported catalysts generally lead to higher space-time yields thanlower contents.

In general, use is made of supported catalysts whose content ofcatalytically active metals is from 0.1 to 90% by weight, preferablyfrom 0.5 to 40% by weight, based on the total catalyst. Since thesecontent figures are based on the total catalyst including supportmaterials but the different support materials have very differentdensities and specific surface areas, it is possible for the contents tobe above or below the ranges specified without this having an adverseeffect on the result of the process of the present invention.

It is also possible for a plurality of catalytically active metals to beapplied to the respective support material.

Furthermore, the catalytically active metals can be applied to thesupport by, for example, the methods described in DE-A 25 19 817, EP-A 0147 219 and EP-A 0 285 420. In the catalysts described in thesedocuments, the catalytically active metals are present as alloys. Theseare produced, for example, by impregnation with a salt or complex of theabovementioned metals and subsequent thermal treatment and/or reduction.

The activation of both the precipitated catalysts and the supportedcatalysts can also be carried out in situ at the beginning of thereaction by means of the hydrogen present, but these catalysts arepreferably activated separately before use.

Support materials which can be used are, in general, the oxides ofaluminum and titanium, zirconium dioxide, silicon dioxide, clayminerals, e.g. montmorillonites, silicates such as magnesium or aluminumsilicates, zeolites such as ZSM-5 or ZSM-10 zeolites, and activatedcarbon. Preferred support materials are aluminum oxides, titaniumdioxide, silicon dioxide, zirconium dioxide and activated carbon.Mixtures of various support materials can also serve as support for thecatalysts to be used in the process of the present invention.

Examples of heterogeneous catalysts which can be used in the process ofthe present invention are: cobalt on activated carbon, cobalt on silicondioxide, cobalt on aluminum oxide, rhenium on activated carbon, rheniumon silicon dioxide, rhenium/tin on activated carbon, rhenium/platinum onactivated carbon, copper on activated carbon, copper/silicon dioxide,copper/aluminum oxide, copper chromite, barium copper chromite,copper/aluminum oxide/manganese oxide, copper/aluminum oxide/zinc oxideand the catalysts described in DE-A 39 32 332, U.S. Pat. No. 3 449 445,EP-A 0 044 444, EP-A 0 147 219, DE-A 39 04 083, DE-A 23 21 101, EP-A 0415 202, DE-A 23 66 264, EP-A 0 552 463 and EP-A 0 100 406.

Preferred catalysts comprise at least one of the metals copper,manganese, rhenium, cobalt, chromium, palladium, platinum or nickel.Particular preference is given to copper, cobalt, palladium, platinum orrhenium.

The ratio of the individual desired products present in thehydrogenation product, viz. 1,4-butanediol, THF and GBL, can vary. Forexample, the molar ratio of the products may be: butanediol 50-95 mol %,THF 2-40 mol %, GBL 0.1-20 mol %, where the sum of the mole fractions of1,4-butanediol, THF and GBL is 100 mol %. The ratio of the desiredproducts is determined predominantly by the parameters pressure,temperature, residence time and catalyst in the hydrogenation. Thus, forexample, the GBL content can be reduced to virtually 0 if thehydrogenation is carried out at a high pressure, a low temperature and along residence time. The THF content can be high when the hydrogenationcatalyst has acid centers.

Further products which are or can be present in the hydrogenationproduct are, for example, water, n-butanol, n-propanol and esters ofsuccinic acid and also the absorption alcohol. The alcohol components ofthe succinic esters can be both the previously used absorption alcoholand butanediol. The succinic esters can be recirculated together withthe absorption alcohol.

The work-up of the reaction products can be carried out in a manner withwhich those skilled in the art are familiar. Preference is given to awork-up by distillation. Here, for example, low boilers such as THF andany water, butanol or propanol present can first be taken off at the topof a distillation column and the remaining products butanediol and GBLcan then be distilled off from the bottoms. The bottoms, which consistpredominantly of the absorption alcohol, are subsequently recirculatedto the MA absorption. If desired, a small bleed stream can bedischarged.

The following examples illustrate the invention.

EXAMPLE 1

A product gas stream from an n-butane oxidation, which comprised about1% by volume of MA and 99% by volume of air, was passed at 100° C. intothe lower part of a packed column which had about 25 theoretical platesand was provided with intermediate coolers to remove the heat ofabsorption. The intermediate coolers were held at about 65° C.1,4-Cyclohexanedimethanol at 70° C. was pumped to the top of the column.A temperature of 87° C. was established at the top of the column and atemperature of 105° C. was established at the bottom of the column.

Overall, about 25 g of MA were bound by 144 g of1,4-cyclohexanedimethanol. The absorption product was subsequentlyheated at 200° C. for 2.5 hours and then at 225° C. for another 30minutes in a residence time vessel. The water of esterification wasdistilled off during this procedure. The product then contained 0.19% byweight of water and the acid number was 19.1 mg KOH/g. To improve thepumpability, the product was diluted with the same amount of ethyleneglycol dimethyl ether and was hydrogenated continuously over 25 ml of aCu catalyst from Süd-Chemie AG, Munich (T 4489) at 220° C. and 220 bar(feed=about 20 g/h, tube reactor, downflow mode without productrecirculation).

The colorless hydrogenation product comprised, excludingcyclohexanedimethanol and ethylene glycol dimethyl ether, 70 mol % of1,4-butanediol, 10 mol % of THF and less than 1 mol % of GBL. Theremainder was predominantly n-butanol and diesters of succinic acid. Theproduct was worked up by distillation, with THF, butanol, ethyleneglycol dimethyl ether, GBL and 1,4-butanediol being distilled off. Thebottoms consisted predominantly of cyclohexanedimethanol and diesters ofsuccinic acid.

EXAMPLE 2

Using a method analogous to Example 1, MA was absorbed in1,6-hexanediol. Here, 73.5 g of MA were reacted with 177 g ofhexanediol. The after-esterification was carried out for 1 hour at 200°C and 225° C. The esterification product contained 0.27% by weight ofwater and had an acid number of 8.3 mg KOH/g. The product was dilutedwith ethylene glycol dimethyl ether as described in Example 1 before thehydrogenation and the hydrogenation was carried out as in Example 1 butthis time at 220 bar or 110 bar. At 220 bar, the colorless hydrogenationproduct comprised, excluding hexanediol and ethylene glycol dimethylether, 85 mol % of 1,4-butanediol, 3 mol % of THF and about 0.2 mol % ofGBL. The remainder was predominantly butanol and diesters of succinicacid.

At 110 bar, the colorless hydrogenation product comprised, excludinghexanediol and ethylene glycol dimethyl ether, 68 mol % of1,4-butanediol, 2 mol % of THF and 8 mol % of GBL. The remainder waspredominantly butanol and diesters of succinic acid.

Comparative Example C1

(DE-A 31 06 819)

Using a method analogous to Example 2, MA was reacted with hexanediol,but the esterification was carried out for 3 hours at 150° C. Theesterification product had a water content of 0.2% and an acid number of49 mg KOH/g. The hydrogenation was carried out at 110 bar as describedin Example 2. After only a short time, the previously colorlesshydrogenation product became reddish brown (copper/manganese) and thehydrogenation activity dropped.

We claim:
 1. A process for preparing 1,4-butanediol, tetrahydrofuran andγ-butyrolactone by oxidation of butane to give a product streamcomprising maleic anhydride, absorption of maleic anhydride from theproduct stream using a high-boiling alcohol to give a liquid absorptionproduct comprising monoesters and diesters of maleic acid and alsohigh-boiling alcohol, after-esterification of the liquid absorptionproduct and subsequent hydrogenation of the after-esterified product inthe liquid phase, wherein the high-boiling alcohol is a polyhydricalcohol having a boiling point at atmospheric pressure of above 233° C.and the after-esterified product has an acid number of less than 30 mgKOH/g and a water content of less than 1% by weight.
 2. A process asclaimed in claim 1, wherein the alcohol used is 1,6-hexanediol or1,4-cyclohexanedimethanol.
 3. A process as claimed in claim 1, whereinthe after-esterification is carried out in the apparatus used for theabsorption of maleic anhydride.
 4. A process as claimed in claim 1,wherein the after-esterification is carried out at from 160° C. to 300°C.
 5. A process as claimed in claim 1, wherein the after-esterificationis carried out at a residence time of not more than 3 hours.
 6. Aprocess as claimed in claim 1, wherein the hydrogenation is carried outat from 70° C. to 350° C. and at a pressure of from 20 to 350 bar.
 7. Aprocess as claimed in claim 1, wherein the hydrogenation is carried outin the presence of a catalyst comprising at least one element of groupsIb, VIb, VIIb and VIIIb of the Periodic Table of the Elements.
 8. Aprocess as claimed in claim 7, wherein the hydrogenation catalystcomprises copper.
 9. A process as claimed in claim 1, wherein the ratioof the individual desired products in the hydrogenation product, viz.1,4-butanediol, tetrahydrofuran and γ-butyrolactone, is 50-95 mol % of1,4-butanediol, 2-40 mol % of tetrahydrofuran and 0.1-20 mol % ofγ-butyrolactone, where the sum of the mole fractions of 1,4-butanediol,tetrahydrofuran and γ-butyrolactone is 100 mol %.
 10. A process asclaimed claim 1, wherein the high-boiling alcohol is returned to theabsorption after the hydrogenation.